U.S. patent number 10,465,131 [Application Number 14/902,068] was granted by the patent office on 2019-11-05 for process for the production of light olefins and aromatics from a hydrocarbon feedstock.
This patent grant is currently assigned to SABIC GLOBAL TECHNOLOGIES B.V., SAUDI BASIC INDUSTRIES CORPORATION. The grantee listed for this patent is SABIC Global Technologies B.V., SAUDI BASIC INDUSTRIES CORPORATION. Invention is credited to Arno Johannes Maria Oprins.
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United States Patent |
10,465,131 |
Oprins |
November 5, 2019 |
Process for the production of light olefins and aromatics from a
hydrocarbon feedstock
Abstract
The present invention relates to a process for increasing the
production of a light olefin hydrocarbon compound from a
hydrocarbon feedstock, comprising the following steps of: (a)
feeding a hydrocarbon feedstock into a reaction area for
ringopening (b) separating reaction products, which are generated
from said reaction area, into an overhead stream and a side stream;
(c) feeding the side stream from (b) to a gasoline hydrocracker
(GHC) unit, (d) separating reaction products of said GHC of step
(c) into an overhead stream, which contains hydrogen, methane,
ethane, and liquefied petroleum gas, and a stream, which contains
aromatic hydrocarbon compounds, and a small amount of hydrogen and
non-aromatic hydrocarbon compounds, (e) feeding the overhead stream
from the gasoline hydrocracker (GHC) unit into a steam cracker
unit.
Inventors: |
Oprins; Arno Johannes Maria
(Maastricht, NL) |
Applicant: |
Name |
City |
State |
Country |
Type |
SAUDI BASIC INDUSTRIES CORPORATION
SABIC Global Technologies B.V. |
Riyadh
Bergen op Zoom |
N/A
N/A |
SA
NL |
|
|
Assignee: |
SAUDI BASIC INDUSTRIES
CORPORATION (Riyadh, SA)
SABIC GLOBAL TECHNOLOGIES B.V. (Bergen op Zoom,
NL)
|
Family
ID: |
48700462 |
Appl.
No.: |
14/902,068 |
Filed: |
June 30, 2014 |
PCT
Filed: |
June 30, 2014 |
PCT No.: |
PCT/EP2014/063851 |
371(c)(1),(2),(4) Date: |
December 30, 2015 |
PCT
Pub. No.: |
WO2015/000843 |
PCT
Pub. Date: |
January 08, 2015 |
Prior Publication Data
|
|
|
|
Document
Identifier |
Publication Date |
|
US 20170009157 A1 |
Jan 12, 2017 |
|
Foreign Application Priority Data
|
|
|
|
|
Jul 2, 2013 [EP] |
|
|
13174767 |
|
Current U.S.
Class: |
1/1 |
Current CPC
Class: |
C10G
67/0445 (20130101); C10G 69/00 (20130101); C10G
65/10 (20130101); C10G 69/06 (20130101); C10G
2400/30 (20130101); Y02P 30/40 (20151101); C10G
2400/20 (20130101); C10G 2300/1059 (20130101); C10G
2300/1044 (20130101) |
Current International
Class: |
C10G
69/06 (20060101); C10G 67/04 (20060101); C10G
69/00 (20060101); C10G 65/10 (20060101) |
References Cited
[Referenced By]
U.S. Patent Documents
Foreign Patent Documents
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|
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|
|
101208412 |
|
Jun 2008 |
|
CN |
|
0192059 |
|
Aug 1986 |
|
EP |
|
2364879 |
|
Apr 1978 |
|
FR |
|
2366239 |
|
Apr 1978 |
|
FR |
|
2162082 |
|
Jan 1986 |
|
GB |
|
2009508881 |
|
Mar 2009 |
|
JP |
|
0244306 |
|
Jun 2002 |
|
WO |
|
2007055488 |
|
May 2007 |
|
WO |
|
2010102712 |
|
Jun 2010 |
|
WO |
|
2012135111 |
|
Oct 2012 |
|
WO |
|
WO 2016/146326 |
|
Sep 2016 |
|
WO |
|
Other References
Halim et al., "Effect of Operating Conditions on
Hydrodesulfurization of Vacuum Gas Oil," Iraqi Journal of Chemical
and Petroleum Engineering, 2008, 57-67. cited by applicant .
Office Action issued in European Application No. 14733659, dated
Feb. 10, 2017. cited by applicant .
Alfke et al. (2007) Oil Refining, Ullmann's Encyclopedia of
Industrial Chemistry. cited by applicant .
English Abstract of WO2010102712(A2); Date of Publication: Sep. 19,
2010; 2 Pages. cited by applicant .
Folkins (2000) Benzene, Ullmann's Encyclopedia of Industrial
Chemistry. cited by applicant .
International Search Report for International Application No.
PCT/EP2014/063851; International Filing Date: Jun. 30, 2014; dated
Sep. 8, 2014; 5 Pages. cited by applicant .
Machine Translation of FR2364879; Date of Publication: Apr. 14,
1978; 26 Pages. cited by applicant .
Machine Translation of FR2366239(A1); Date of Publication: Apr. 28,
1978; 12 Pages. cited by applicant .
Speight (2005) Petroleum Refinery Process, Kirk-Othmer Encyclopedia
of Chemical Technology. cited by applicant .
Table VI, p. 295, Pyrolysis: Theory and Industrial Practice by Lyle
R Albright et al., Academic Press 1983. cited by applicant .
Written Opinion of the International Searching Authority for
International Application No. PCT/EP2014/063851; International
Filing Date: Jun. 30, 2014; dated Sep. 8, 2014; 6 Pages. cited by
applicant .
Office Action issued in Chinese Patent Application No.
201480037272.7 dated Sep. 25, 2017. cited by applicant .
Examination Report issued in Gulf Cooperation Coucil Application
No. 2014/27470, dated Mar. 29, 2017. cited by applicant .
Office Action issued in Japanese Patent Application No.
2016-522561, dated Feb. 20, 2018. cited by applicant.
|
Primary Examiner: Singh; Prem C
Assistant Examiner: Doyle; Brandi M
Attorney, Agent or Firm: Norton Rose Fulbright US LLP
Claims
The invention claimed is:
1. A process for increasing the production of a light olefin
hydrocarbon compound from a hydrocarbon feedstock, consisting of
the steps of: feeding a hydrocarbon feedstock into a reaction area
for ring opening, wherein the process conditions prevailing in said
reaction area for ring opening are a temperature from 300.degree.
C. to 500.degree. C. and a pressure from 2 to 10 MPa together with
from 100 to 300 kg of hydrogen per 1,000 kg of feedstock over an
aromatic hydrogenation catalyst, wherein said aromatic
hydrogenation catalyst comprises from 0.0001 to 5 weight % of one
or more metals selected from the group consisting of Ni, W, and Mo,
passing the resulting stream to a ring cleavage unit at a
temperature from 200.degree. C. to 600.degree. C. and a pressure
from 1 to 12 MPa together with from 50 to 200 kg of hydrogen per
1,000 kg of said resulting stream over a ring cleavage catalyst
comprising from 0.0001 to 5 weight % of one or more metals selected
from the group consisting of Pd, Ru, Ir, Os, Cu, Co, Ni, Pt, Fe,
Zn, Ga, In, Mo, W, and V on a support selected from the group of
synthetic zeolites having the characteristics of ZSM-5, ZSM-11,
ZSM-12, ZSM-23, Beta and MCM-22; separating reaction products,
which are generated from said reaction area, into an overhead
stream and a side stream; feeding the side stream from (b) to a
gasoline hydrocracker (GHC) unit operating at a temperature range
of 400-580.degree. C., a Weight Hourly Space Velocity (WHSV) of
0.1-10 h-1 and a pressure range of 0.3-5 MPa, wherein said gasoline
hydrocracker (GHC) unit is operated at a temperature higher than
said ring opening reaction area, and wherein said gasoline
hydrocracker (GHC) unit is operated at a pressure lower than said
ring opening reaction area, separating reaction products of said
GHC of step (c) into an overhead gas stream, comprising C2-C4
paraffins, hydrogen and methane and a stream comprising aromatic
hydrocarbon compounds and non-aromatic hydrocarbon compounds,
feeding the overhead gas stream from the gasoline hydrocracker
(GHC) unit into a steam cracker unit; separating reaction products
of said steam cracking unit into an overhead stream, comprising
C2-C6 alkanes, a middle stream comprising C2-olefins, C3-olefins
and C4-olefins, and a first bottom stream comprising predominantly
carbon black oil (CBO) and cracked distillates (CD), and a second
bottom stream comprising aromatic hydrocarbon compounds and
non-aromatic hydrocarbon compounds; and feeding said first bottom
stream into said reaction area for ring opening; wherein the
hydrocarbon feedstock consists of a fraction of at least one member
selected from the group consisting of a conventional petroleum
having an API gravity of more than 20.degree. API as measured by
the ASTM D287 standard and a light crude oil having an API gravity
of more than 30.degree. API.
Description
CROSS REFERENCE TO RELATED APPLICATIONS
This application is a National Stage application of
PCT/EP2014/063851, filed Jun. 30, 2014, which claims the benefit of
European Application No. 13174767.7, filed Jul. 2, 2013, both of
which are incorporated by reference in their entirety herein.
The present invention relates to a process for the production of
light olefins and aromatics from a hydrocarbon feedstock.
Conventionally, crude oil is processed, via distillation, into a
number of cuts such as naphtha, gas oils and residua. Each of these
cuts has a number of potential uses such as for producing
transportation fuels such as gasoline, diesel and kerosene or as
feeds to some petrochemicals and other processing units.
Light crude oil cuts such as naphthas and some gas oils can be used
for producing light olefins and single ring aromatic compounds via
processes such as steam cracking in which the hydrocarbon feed
stream is evaporated and diluted with steam then exposed to a very
high temperature (800.degree. C. to 860.degree. C.) in short
residence time (<1 second) furnace (reactor) tubes. In such a
process the hydrocarbon molecules in the feed are transformed into
(on average) shorter molecules and molecules with lower hydrogen to
carbon ratios (such as olefins) when compared to the feed
molecules. This process also generates hydrogen as a useful
by-product and significant quantities of lower value co-products
such as methane and C9+ Aromatics and condensed aromatic species
(containing two or more aromatic rings which share edges).
Typically, the heavier (or higher boiling point) aromatic species,
such as residua are further processed in a crude oil refinery to
maximize the yields of lighter (distillable) products from the
crude oil. This processing can be carried out by processes such as
hydro-cracking (whereby the hydro-cracker feed is exposed to a
suitable catalyst under conditions which result in some fraction of
the feed molecules being broken into shorter hydrocarbon molecules
with the simultaneous addition of hydrogen). Heavy refinery stream
hydrocracking is typically carried out at high pressures and
temperatures and thus has a high capital cost.
An aspect of such a combination of crude oil distillation and steam
cracking of the lighter distillation cuts is the capital and other
costs associated with the fractional distillation of crude oil.
Heavier crude oil cuts (i.e. those boiling beyond
.about.350.degree. C.) are relatively rich in substituted aromatic
species and especially substituted condensed aromatic species
(containing two or more aromatic rings which share edges) and under
steam cracking conditions these materials yield substantial
quantities of heavy by products such as C9+ aromatics and condensed
aromatics. Hence, a consequence of the conventional combination of
crude oil distillation and steam cracking is that a substantial
fraction of the crude oil is not processed via the steam cracker as
the cracking yield of valuable products from heavier cuts is not
considered to be sufficiently high, or at least when compared to
alternative refinery value.
Another aspect of the technology discussed above is that even when
only light crude oil cuts (such as naphtha) are processed via steam
cracking a significant fraction of the feed stream is converted
into low value heavy by-products such as C9+ aromatics and
condensed aromatics. With typical naphthas and gas oils these heavy
by-products might constitute 2 to 25% of the total product yield
(Table VI, Page 295, Pyrolysis: Theory and Industrial Practice by
Lyle F. Albright et al, Academic Press, 1983). Whilst this
represents a significant financial downgrade of expensive naphtha
in lower value material on the scale of a conventional steam the
yield of these heavy by-products to does not typically justify the
capital investment required to up-grade these materials (e.g. by
hydrocracking) into streams that might produce significant
quantities of higher value chemicals. This is partly because
hydrocracking plants have high capital costs and, as with most
petrochemicals processes, the capital cost of these units typically
scales with throughput raised to the power of 0.6 or 0.7.
Consequently, the capital costs of a small scale hydro-cracking
unit are normally considered to be too high to justify such an
investment to process steam cracker heavy by-products.
Another aspect of the conventional hydrocracking of heavy refinery
streams such as residua is that this is typically carried out under
compromise conditions chosen to achieve the desired overall
conversion. As the feed streams contain a mixture of species with a
range of ease of cracking this result in some fraction of the
distillable products formed by hydrocracking of relatively easily
hydrocracked species being further converted under the conditions
necessary to hydrocrack species more difficult to hydrocrack. This
increases the hydrogen consumption and heat management difficulties
associated with the process and also increases the yield of light
molecules such as methane at the expense of more valuable
species.
A result of such a combination of crude oil distillation and steam
cracking of the lighter distillation cuts is that steam cracking
furnace tubes are typically unsuitable for the processing of cuts
which contain significant quantities of material with a boiling
point greater than .about.350.degree. C. as it is difficult to
ensure complete evaporation of these cuts prior to exposing the
mixed hydrocarbon and steam stream to the high temperatures
required to promote thermal cracking. If droplets of liquid
hydrocarbon are present in the hot sections of cracking tubes coke
is rapidly deposited on the tube surface which reduces heat
transfer and increases pressure drop and ultimately curtails the
operation of the cracking tube necessitating a shut-down of the
tube to allow for decoking. Due to this difficulty a significant
proportion of the original crude oil cannot be processed into light
olefins and aromatic species via a steam cracker.
The LCO Unicracking process of UOP uses partial conversion
hydrocracking to produce high quality gasoline and diesel stocks in
a simple once-through flow scheme. The feedstock is processed over
a pretreatment catalyst and then hydrocracked in the same stage.
The products are subsequently separated without the need for liquid
recycle. The LCO Unicracking process can be designed for lower
pressure operation, that is the pressure requirement will be
somewhat higher than high severity hydrotreating but significantly
lower than a conventional partial conversion and full conversion
hydrocracking unit design. The upgraded middle distillate product
makes a suitable ultra-low sulfur diesel (ULSD) blending component.
The naphtha product from low-pressure hydrocracking of LCO has
ultra-low sulfur and high octane and can be directly blended into
the ultra-low sulfur gasoline (ULSG) pool.
U.S. Pat. No. 7,513,988 relates to a process to treat compounds
comprising two or more fused aromatic rings to saturate at least
one ring and then cleave the resulting saturated ring from the
aromatic portion of the compound to produce a C2-4 alkane stream
and an aromatic stream. Such a process may be integrated with a
hydrocarbon (e.g. ethylene) (steam) cracker so that hydrogen from
the cracker may be used to saturate and cleave the compounds
comprising two or more aromatic rings and the C2-4 alkane stream
may be fed to the hydrocarbon cracker, or may be integrated with a
hydrocarbon cracker (e.g. steam cracker) and an ethylbenzene unit,
that is to treat the heavy residues from processing oil sands, tar
sands, shale oils or any oil having a high content of fused ring
aromatic compounds to produce a stream suitable for petrochemical
production.
US2005/0101814 relates to a process for improving the paraffin
content of a feedstock to a steam cracking unit, comprising:
passing a feedstream comprising C5 through C9 hydrocarbons
including C5 through C9 normal paraffins into a ring opening
reactor, the ring opening reactor comprising a catalyst operated at
conditions to convert aromatic hydrocarbons to naphtenes and a
catalyst operated at conditions to convert naphtenes to paraffins,
and producing a second feedstream; and passing at least a portion
of the second feedstream to a steam cracking unit.
U.S. Pat. No. 7,067,448 relates to a process for the manufacture of
n-alkanes from mineral oil fractions and fractions from thermal or
catalytic conversion plants containing cyclic alkanes, alkenes,
cyclic alkenes and/or aromatic compounds. More in detail, this
publication refers to a process for processing mineral oil
fractions rich in aromatic compounds, in which the cyclic alkanes
obtained after the hydrogenation of the aromatic compounds are
converted to n-alkanes of a chain length which as far as possible
is less than that of the charged carbons.
US2009/173665 relates to a catalyst and process for increasing the
monoaromatics content of hydrocarbon feedstocks that include
polynuclear aromatics, wherein the increase in monoaromatics can be
achieved with an increase in gasoline/diesel yields and while
reducing unwanted compounds thereby providing a route for upgrading
hydrocarbons that include significant quantities of polynuclear
aromatics.
U.S. Pat. No. 4,137,147 (corresponding to FR 2 364 879 and FR 2 366
239) relates to a selective process for producing light olefinic
hydrocarbons chiefly those with 2 and 3 carbon atoms respectively
per molecule, particularly ethylene and propylene, which are
obtained by hydrogenolysis or hydrocracking followed with
steam-cracking.
U.S. Pat. No. 3,842,138 relates to a method of thermal cracking in
the presence of hydrogen of a charge of hydrocarbons of petroleum
wherein the hydrocracking process is carried out under a pressure
of 5 and 70 bars at the outlet of the reactor with very short
residence times of 0.01 and 0.5 second and a temperature range at
the outlet of the reactor extending from 625 to 1000.degree. C.
The LCO-process as discussed above relates to full conversion
hydrocracking of LCO to naphtha, in which LCO is a mono-aromatics
and di-aromatics containing stream. A consequence of the full
conversion hydrocracking is that a highly naphthenic, low octane
naphtha is obtained that must be reformed to produce the octane
required for product blending.
An object of the present invention is to provide a method for
upgrading naphtha, naphtha condensates and heavy tail feeds to
aromatics and LPG cracker feeds.
Another object of the present invention is to provide a process for
the production of light olefins and aromatics from a hydrocarbon
feedstock in which a high yield of ethylene and propylene can be
attained.
Another object of the present invention is to provide a process for
the production of light olefins and aromatics from a hydrocarbon
feedstock in which a broad spectrum of hydrocarbon feedstocks can
be processed, i.e. a high feed flexibility.
Another object of the present invention is to provide a process for
the production of light olefins and aromatics from a hydrocarbon
feedstock in which a high yield of aromatics can be attained.
The present invention relates to a process for increasing the
production of a light olefin hydrocarbon compound from a
hydrocarbon feedstock, comprising the following steps of:
(a) feeding a hydrocarbon feedstock into a reaction area for
ringopening
(b) separating reaction products, which are generated from said
reaction area, into an overhead stream and a side stream;
(c) feeding the side stream from (b) to a gasoline hydrocracker
(GHC) unit,
(d) separating reaction products of said GHC of step (c) into an
overhead gas stream, comprising C2-C4 paraffins, hydrogen and
methane and a stream comprising aromatic hydrocarbon compounds and
non-aromatic hydrocarbon compounds,
(e) feeding the overhead gas stream from the gasoline hydrocracker
(GHC) unit into a steam cracker unit.
On basis of these steps (a)-(e) one or more of the present objects
can be attained. The present inventors found that a full conversion
hydrocracking step can be used resulting in the direct conversion
of the produced naphthenic naphtha in the GHC into a high quality
BTX stream and a very good LPG cracker feed. A difference to the
LCO-X process being that in addition to retaining the aromatic
rings (or the last aromatic ring in case of ring opening of di and
tri aromatics) the present invention also converts the naphthenic
species in a single step process largely into BTX as well as a
result of the specific conditions in the GHC without the need for a
reformer that doesn't yield the high value LPG product suitable for
steam cracking, nor the direct production of high quality BTX. When
using a ring opening step the advantage of the GHC processing step
is that the aromatic product obtained is upgraded to produce BTX
and LPG rather than a mix of higher mono-aromatics that would not
be suitable for chemicals production but only has a value in
gasoline blending. The present invention is more focusing on using
the GHC platform to directly produce high quality BTX and high
value LPG cracker feed, i.e. also upgrading the `side chains` of
the higher mono-aromatics produced in the LCO-process. Moreover,
according to the present invention paraffin species will be
converted into high value LPG (and hydrogen `re-claimed` further
downstream) and naphthenic species will be converted into BTX, that
is to retain the aromatic rings and break down poly aromatic
components retaining the last aromatic ring. Effectively the
combination of process steps in the present invention allows the
present inventors to control the amount of LPG leading to light
olefins versus the amount of BTX obtained.
The preferred process conditions for the reaction area for
ringopening comprise passing said feed stream to a ring saturation
unit at a temperature from 300 [deg.] C. to 500 [deg.] C. and a
pressure from 2 to 10 MPa together with from 100 to 300 kg of
hydrogen per 1,000 kg of feedstock over an aromatic hydrogenation
catalyst and passing the resulting stream to a ring cleavage unit
at a temperature from 200 [deg.] C. to 600 [deg.] C. and a pressure
from 1 to 12 MPa together with from 50 to 200 kg of hydrogen per
1,000 kg of said resulting stream over a ring cleavage catalyst.
The resulting product can be separated into a C2-4 alkanes stream,
a liquid paraffinic stream and an aromatic stream. The aromatic
hydrogenation catalyst comprises from 0.0001 to 5 weight % of one
or more metals selected from the group consisting of Ni, W, and Mo.
The ring cleavage catalyst comprises from 0.0001 to 5 weight % of
one or more metals selected from the group consisting of Pd, Ru,
Is, Os, Cu, Co, Ni, Pt, Fe, Zn, Ga, In, Mo, W, and V on a support,
i.e. the support is selected from the group of synthetic zeolites
having the characteristics of ZSM-5, ZSM-11, ZSM-12, ZSM-23, Beta
and MCM-22.
In a preferred embodiment the present method further comprises
pretreating the hydrocarbon feedstock in an aromatics extraction
unit, from which aromatics extraction unit its bottom stream is fed
into said reaction area for ringopening and its overhead stream is
fed into said steam cracker unit.
The aromatics extraction unit is chosen from the group of the type
of a distillation unit, a solvent extraction unit and a molecular
sieve, or even a combination thereof.
In the embodiment of a solvent extraction unit its overhead stream
is washed for removal of solvent, wherein the thus recovered
solvent is returned into said solvent extraction unit and the
overhead stream thus washed being fed into said steam cracker unit.
In such an extraction unit the liquid hydrocarbon feed is in the
solvent extraction step first contacted with an immiscible solvent
selective for aromatics separation in a suitable solvent extraction
column. The boiling temperature of the immiscible solvents
selective for aromatics separation must be higher than the boiling
temperature of the components to be separated, i.e. extract
containing aromatics and naphthenes. A preferred temperature
difference between immiscible solvent and the extract is in the
range of 10 to 20 degr Celsius. In addition, the immiscible solvent
must not decompose at the applied temperatures, i.e. the immiscible
solvent must be temperature stabile at the specific process
temperature. Examples of solvents are sulfolane, tetra ethylene
glycol or N-Methyl pyrolidone. These species are often used in
combination with other solvents or other chemicals (sometimes
called co-solvents) such as water and/or alcohols. To minimize the
risk of damaging the hydrocracking catalyst in the present process,
it is preferred to use a non-nitrogen containing solvent such as
sulfolane. As the solvent (even when it contains significant
quantities of dissolved hydrocarbons) has a higher density than the
hydrocarbon species it tends to separate to the base of the
extraction column and is withdrawn from there. This "rich solvent"
(i.e. solvent containing dissolved hydrocarbons) contains aromatic
species which were present in the feed liquid as well as other
species which are somewhat soluble in the solvent such as light
paraffins, naphthenic species as well as some of the organo-sulphur
species present in the feed. With conventional technologies the
presence of the non-aromatic hydrocarbon species causes a
difficulty which requires these species to be stripped from the
"rich solvent" in a distillation column (together with some of the
lower boiling point aromatic compounds) and returned to the solvent
extraction column. To ensure that the aromatic product stream is
essentially free from non-aromatics contaminants it is necessary to
expend significant quantities of energy in stripping out even minor
traces of these species from the solvent.
According to a preferred embodiment the reaction products of said
steam cracking unit are separated into an overhead stream,
comprising C2-C6 alkanes, a middle stream, comprising C2-olefins,
C3-olefins and C4-olefins, and a first bottom stream, comprising
C9+ hydrocarbons, and a second bottom stream comprising aromatic
hydrocarbon compounds and non-aromatic hydrocarbon compounds.
The present process further comprises returning said overhead
stream to said steam cracking unit.
In a preferred embodiment of the present invention the second
bottom stream is fed into said gasoline hydrocracker (GHC) unit. It
is also preferred to feed the first bottom stream predominantly
containing carbon black oil (CBO) and cracked distillates (CD) into
said reaction area for ringopening. According to another embodiment
the bottom stream from reaction products of said gasoline
hydrocracker (GHC) unit is separated in a BTX rich fraction and in
heavy fraction.
The overhead stream from the gasoline hydrocracker (GHC) unit is
preferably fed into a dehydrogenation unit, especially the C3-C4
fraction thereof. In addition it is also preferred to feed said
overhead stream from the reaction area for ring opening to a
dehydrogenation unit, especially the C3-C4 fraction thereof.
According to the present invention the LPG rich fractions can thus
be sent either to the steam cracker unit and/or to the
dehydrogenation unit. This provides a high level of flexibility and
product diversity. The overhead stream from the reaction area for
ring opening and the overhead gas stream from the gasoline
hydrocracker (GHC) unit can be indicated as LPG rich fractions.
Processes for the dehydrogenation of lower alkanes such as propane
and butanes are described as lower alkane dehydrogenation process.
The term "propane dehydrogenation unit" relates to a petrochemical
process unit wherein a propane feedstream is converted into a
product comprising propylene and hydrogen. Accordingly, the term
"butane dehydrogenation unit" relates to a process unit for
converting a butane feedstream into C4 olefins.
The present process further comprises recovering a stream rich in
mono aromatics from said hydrocarbon feedstock of step (a) and
feeding the stream thus recovered to said gasoline hydrocracker
(GHC) unit, and recovering a stream rich in mono aromatics from
said bottom stream of said aromatics extraction unit and feeding
the stream thus recovered to said gasoline hydrocracker (GHC)
unit.
From a hydrogen consumption perspective it is preferred to recover
hydrogen from the reaction products of said steam cracking unit and
feeding the hydrogen thus recovered to said gasoline hydrocracker
(GHC) unit and/or said reaction area for ring opening, especially
to recover hydrogen from said dehydrogenation unit and feeding the
hydrogen thus recovered to said gasoline hydrocracker (GHC) unit
and/or said reaction area for ring opening.
Examples of preferred hydrocarbon feedstock to be fed into said
reaction area for ring opening are chosen from the group of gasoil,
vacuum gas oil (VGO), naphtha and pretreated naphtha, or a
combination thereof.
The process conditions prevailing in said reaction area for ring
opening have been mentioned above.
The process conditions prevailing in said gasoline hydrocracker
(GHC) unit comprise a temperature of 300-450.degree. C., a pressure
of 300-5000 kPa gauge and a Weight Hourly Space Velocity of 0.1-10
h-1, preferably a temperature of 300-400.degree. C., a pressure of
600-3000 kPa gauge and a Weight Hourly Space Velocity of 0.2-2
h-1.
The process conditions prevailing in said steam cracking unit will
be discussed hereafter.
In the present process preferred examples of feedstock to be sent
directly into said steam cracking unit comprise a hydrocarbon
feedstock not treated in a series of reaction area(s) for
ringopening and gasoline hydrocracker (GHC) unit(s)
The present invention further relates to the use of a gaseous light
fraction of a multi stage ring opened hydrocracked hydrocarbon
feedstock as a feedstock for a steam cracking unit.
The term "crude oil" as used herein refers to the petroleum
extracted from geologic formations in its unrefined form. Any crude
oil is suitable as the source material for the process of this
invention, including Arabian Heavy, Arabian Light, other Gulf
crudes, Brent, North Sea crudes, North and West African crudes,
Indonesian, Chinese crudes and mixtures thereof, but also shale
oil, tar sands and bio-based oils. The crude oil is preferably
conventional petroleum having an API gravity of more than
20.degree. API as measured by the ASTM D287 standard. More
preferably, the crude oil used is a light crude oil having an API
gravity of more than 30.degree. API. Most preferably, the crude oil
comprises Arabian Light Crude Oil. Arabian Light Crude Oil
typically has an API gravity of between 32-36.degree. API and a
sulfur content of between 1.5-4.5 wt-%.
The term "petrochemicals" or "petrochemical products" as used
herein relates to chemical products derived from crude oil that are
not used as fuels. Petrochemical products include olefins and
aromatics that are used as a basic feedstock for producing
chemicals and polymers. High-value petrochemicals include olefins
and aromatics. Typical high-value olefins include, but are not
limited to, ethylene, propylene, butadiene, butylene-1,
isobutylene, isoprene, cyclopentadiene and styrene. Typical
high-value aromatics include, but are not limited to, benzene,
toluene, xylene and ethyl benzene.
The term "fuels" as used herein relates to crude oil-derived
products used as energy carrier. Unlike petrochemicals, which are a
collection of well-defined compounds, fuels typically are complex
mixtures of different hydrocarbon compounds. Fuels commonly
produced by oil refineries include, but are not limited to,
gasoline, jet fuel, diesel fuel, heavy fuel oil and petroleum
coke.
The term "gases produced by the crude distillation unit" or "gases
fraction" as used herein refers to the fraction obtained in a crude
oil distillation process that is gaseous at ambient temperatures.
Accordingly, the "gases fraction" derived by crude distillation
mainly comprises C1-C4 hydrocarbons and may further comprise
impurities such as hydrogen sulfide and carbon dioxide. In this
specification, other petroleum fractions obtained by crude oil
distillation are referred to as "naphtha", "kerosene", "gasoil" and
"resid". The terms naphtha, kerosene, gasoil and resid are used
herein having their generally accepted meaning in the field of
petroleum refinery processes; see Alfke et al. (2007) Oil Refining,
Ullmann's Encyclopedia of Industrial Chemistry and Speight (2005)
Petroleum Refinery Processes, Kirk-Othmer Encyclopedia of Chemical
Technology. In this respect, it is to be noted that there may be
overlap between the different crude oil distillation fractions due
to the complex mixture of the hydrocarbon compounds comprised in
the crude oil and the technical limits to the crude oil
distillation process. Preferably, the term "naphtha" as used herein
relates to the petroleum fraction obtained by crude oil
distillation having a boiling point range of about 20-200.degree.
C., more preferably of about 30-190.degree. C. Preferably, light
naphtha is the fraction having a boiling point range of about
20-100.degree. C., more preferably of about 30-90.degree. C. Heavy
naphtha preferably has a boiling point range of about
80-200.degree. C., more preferably of about 90-190.degree. C.
Preferably, the term "kerosene" as used herein relates to the
petroleum fraction obtained by crude oil distillation having a
boiling point range of about 180-270.degree. C., more preferably of
about 190-260.degree. C. Preferably, the term "gasoil" as used
herein relates to the petroleum fraction obtained by crude oil
distillation having a boiling point range of about 250-360.degree.
C., more preferably of about 260-350.degree. C. Preferably, the
term "resid" as used herein relates to the petroleum fraction
obtained by crude oil distillation having a boiling point of more
than about 340.degree. C., more preferably of more than about
350.degree. C.
The term "aromatic hydrocarbons" or "aromatics" is very well known
in the art. Accordingly, the term "aromatic hydrocarbon" relates to
cyclically conjugated hydrocarbon with a stability (due to
delocalization) that is significantly greater than that of a
hypothetical localized structure (e.g. Kekule structure). The most
common method for determining aromaticity of a given hydrocarbon is
the observation of diatropicity in the 1H NMR spectrum, for example
the presence of chemical shifts in the range of from 7.2 to 7.3 ppm
for benzene ring protons.
The terms "naphthenic hydrocarbons" or "naphthenes" or
"cycloalkanes" is used herein having its established meaning and
accordingly relates types of alkanes that have one or more rings of
carbon atoms in the chemical structure of their molecules.
The term "olefin" is used herein having its well-established
meaning. Accordingly, olefin relates to an unsaturated hydrocarbon
compound containing at least one carbon-carbon double bond.
Preferably, the term "olefins" relates to a mixture comprising two
or more of ethylene, propylene, butadiene, butylene-1, isobutylene,
isoprene and cyclopentadiene.
The term "LPG" as used herein refers to the well-established
acronym for the term "liquefied petroleum gas". LPG generally
consists of a blend of C2-C4 hydrocarbons i.e. a mixture of C2, C3,
and C4 hydrocarbons.
The term "BTX" as used herein relates to a mixture of benzene,
toluene and xylenes.
As used herein, the term "C# hydrocarbons", wherein "#" is a
positive integer, is meant to describe all hydrocarbons having #
carbon atoms. Moreover, the term "C#+ hydrocarbons" is meant to
describe all hydrocarbon molecules having # or more carbon atoms.
Accordingly, the term "C5+ hydrocarbons" is meant to describe a
mixture of hydrocarbons having 5 or more carbon atoms. The term
"C5+ alkanes" accordingly relates to alkanes having 5 or more
carbon atoms.
As used herein, the term "hydrocracker unit" or "hydrocracker"
relates to a refinery unit in which a hydrocracking process is
performed i.e. a catalytic cracking process assisted by the
presence of an elevated partial pressure of hydrogen; see e.g.
Alfke et al. (2007) loc.cit. The products of this process are
saturated hydrocarbons and, depending on the reaction conditions
such as temperature, pressure and space velocity and catalyst
activity, aromatic hydrocarbons including BTX. The process
conditions used for hydrocracking generally includes a process
temperature of 200-600.degree. C., elevated pressures of 0.2-20
MPa, space velocities between 0.1-10 h-1.
Hydrocracking reactions proceed through a bifunctional mechanism
which requires a acid function, which provides for the cracking and
isomerization and which provides breaking and/or rearrangement of
the carbon-carbon bonds comprised in the hydrocarbon compounds
comprised in the feed, and a hydrogenation function. Many catalysts
used for the hydrocracking process are formed by composting various
transition metals, or metal sulfides with the solid support such as
alumina, silica, alumina-silica, magnesia and zeolites.
As used herein, the term "gasoline hydrocracking unit" or "GHC"
refers to a refinery unit for performing a hydrocracking process
suitable for converting a complex hydrocarbon feed that is
relatively rich in aromatic hydrocarbon compounds--such as refinery
unit-derived light-distillate including, but not limited to,
reformer gasoline, FCC gasoline and pyrolysis gasoline (pygas)--to
LPG and BTX, wherein said process is optimized to keep one aromatic
ring intact of the aromatics comprised in the GHC feedstream, but
to remove most of the side-chains from said aromatic ring.
Accordingly, the main product produced by gasoline hydrocracking is
BTX and the process can be optimized to provide chemicals-grade
BTX. Preferably, the hydrocarbon feed that is subject to gasoline
hydrocracking comprises refinery unit-derived light-distillate.
More preferably, the hydrocarbon feed that is subjected to gasoline
hydrocracking preferably does not comprise more than 1 wt-% of
hydrocarbons having more than one aromatic ring. Preferably, the
gasoline hydrocracking conditions include a temperature of
300-580.degree. C., more preferably of 450-580.degree. C. and even
more preferably of 470-550.degree. C. Lower temperatures must be
avoided since hydrogenation of the aromatic ring becomes favorable.
However, in case the catalyst comprises a further element that
reduces the hydrogenation activity of the catalyst, such as tin,
lead or bismuth, lower temperatures may be selected for gasoline
hydrocracking; see e.g. WO 02/44306 A1 and WO 2007/055488. In case
the reaction temperature is too high, the yield of LPG's
(especially propane and butanes) declines and the yield of methane
rises. As the catalyst activity may decline over the lifetime of
the catalyst, it is advantageous to increase the reactor
temperature gradually over the life time of the catalyst to
maintain the hydrocracking conversion rate. This means that the
optimum temperature at the start of an operating cycle preferably
is at the lower end of the hydrocracking temperature range. The
optimum reactor temperature will rise as the catalyst deactivates
so that at the end of a cycle (shortly before the catalyst is
replaced or regenerated) the temperature preferably is selected at
the higher end of the hydrocracking temperature range.
Preferably, the gasoline hydrocracking of a hydrocarbon feedstream
is performed at a pressure of 0.3-5 MPa gauge, more preferably at a
pressure of 0.6-3 MPa gauge, particularly preferably at a pressure
of 1-2 MPa gauge and most preferably at a pressure of 1.2-1.6 MPa
gauge. By increasing reactor pressure, conversion of C5+
non-aromatics can be increased, but this also increases the yield
of methane and the hydrogenation of aromatic rings to cyclohexane
species which can be cracked to LPG species. This results in a
reduction in aromatic yield as the pressure is increased and, as
some cyclohexane and its isomer methylcyclopentane, are not fully
hydrocracked, there is an optimum in the purity of the resultant
benzene at a pressure of 1.2-1.6 MPa.
Preferably, gasoline hydrocracking of a hydrocarbon feedstream is
performed at a Weight Hourly Space Velocity (WHSV) of 0.1-10 h-1,
more preferably at a Weight Hourly Space Velocity of 0.2-6 h-1 and
most preferably at a Weight Hourly Space Velocity of 0.4-2 h-1.
When the space velocity is too high, not all BTX co-boiling
paraffin components are hydrocracked, so it will not be possible to
achieve BTX specification by simple distillation of the reactor
product. At too low space velocity the yield of methane rises at
the expense of propane and butane. By selecting the optimal Weight
Hourly Space Velocity, it was surprisingly found that sufficiently
complete reaction of the benzene co-boilers is achieved to produce
on spec BTX without the need for a liquid recycle.
Accordingly, preferred gasoline hydrocracking conditions thus
include a temperature of 450-580.degree. C., a pressure of 0.3-5
MPa gauge and a Weight Hourly Space Velocity of 0.1-10 h-1. More
preferred gasoline hydrocracking conditions include a temperature
of 470-550.degree. C., a pressure of 0.6-3 MPa gauge and a Weight
Hourly Space Velocity of 0.2-6 h-1. Particularly preferred gasoline
hydrocracking conditions include a temperature of 470-550.degree.
C., a pressure of 1-2 MPa gauge and a Weight Hourly Space Velocity
of 0.4-2 h-1.
The "aromatic ring opening unit" refers to a refinery unit wherein
the aromatic ring opening process is performed. Aromatic ring
opening is a specific hydrocracking process that is particularly
suitable for converting a feed that is relatively rich in aromatic
hydrocarbon having a boiling point in the kerosene and gasoil
boiling point range to produce LPG and, depending on the process
conditions, a light-distillate (ARO-derived gasoline). Such an
aromatic ring opening process (ARO process) is for instance
described in U.S. Pat. Nos. 3,256,176 and 4,789,457. Such processes
may comprise of either a single fixed bed catalytic reactor or two
such reactors in series together with one or more fractionation
units to separate desired products from unconverted material and
may also incorporate the ability to recycle unconverted material to
one or both of the reactors. Reactors may be operated at a
temperature of 200-600.degree. C., preferably 300-400.degree. C., a
pressure of 3-35 MPa, preferably 5 to 20 MPa together with 5-20
wt-% of hydrogen (in relation to the hydrocarbon feedstock),
wherein said hydrogen may flow co-current with the hydrocarbon
feedstock or counter current to the direction of flow of the
hydrocarbon feedstock, in the presence of a dual functional
catalyst active for both hydrogenation-dehydrogenation and ring
cleavage, wherein said aromatic ring saturation and ring cleavage
may be performed. Catalysts used in such processes comprise one or
more elements selected from the group consisting of Pd, Rh, Ru, Ir,
Os, Cu, Co, Ni, Pt, Fe, Zn, Ga, In, Mo, W and V in metallic or
metal sulphide form supported on an acidic solid such as alumina,
silica, alumina-silica and zeolites. In this respect, it is to be
noted that the term "supported on" as used herein includes any
conventional way to provide a catalyst which combines one or more
elements with a catalytic support. A further aromatic ring opening
process (ARO process) is described in U.S. Pat. No. 7,513,988.
Accordingly, the ARO process may comprise aromatic ring saturation
at a temperature of 100-500.degree. C., preferably 200-500.degree.
C. and more preferably 300-500.degree. C., a pressure of 2-10 MPa
together with 5-30 wt-%, preferably 10-30 wt-% of hydrogen (in
relation to the hydrocarbon feedstock) in the presence of an
aromatic hydrogenation catalyst and ring cleavage at a temperature
of 200-600.degree. C., preferably 300-400.degree. C., a pressure of
1-12 MPa together with 5-20 wt-% of hydrogen (in relation to the
hydrocarbon feedstock) in the presence of a ring cleavage catalyst,
wherein said aromatic ring saturation and ring cleavage may be
performed in one reactor or in two consecutive reactors. The
aromatic hydrogenation catalyst may be a conventional
hydrogenation/hydrotreating catalyst such as a catalyst comprising
a mixture of Ni, W and Mo on a refractory support, typically
alumina. The ring cleavage catalyst comprises a transition metal or
metal sulphide component and a support. Preferably the catalyst
comprises one or more elements selected from the group consisting
of Pd, Rh, Ru, Ir, Os, Cu, Co, Ni, Pt, Fe, Zn, Ga, In, Mo, W and V
in metallic or metal sulphide form supported on an acidic solid
such as alumina, silica, alumina-silica and zeolites. By adapting
either single or in combination the catalyst composition, operating
temperature, operating space velocity and/or hydrogen partial
pressure, the process can be steered towards full saturation and
subsequent cleavage of all rings or towards keeping one aromatic
ring unsaturated and subsequent cleavage of all but one ring. In
the latter case, the ARO process produces a light-distillate
("ARO-gasoline") which is relatively rich in hydrocarbon compounds
having one aromatic ring.
As used herein, the term "dearomatization unit" relates to a
refinery unit for the separation of aromatic hydrocarbons, such as
BTX, from a mixed hydrocarbon feed. Such dearomatization processes
are described in Folkins (2000) Benzene, Ullmann's Encyclopedia of
Industrial Chemistry. Accordingly, processes exist to separate a
mixed hydrocarbon stream into a first stream that is enriched for
aromatics and a second stream that is enriched for paraffins and
naphthenes. A preferred method to separate aromatic hydrocarbons
from a mixture of aromatic and aliphatic hydrocarbons is solvent
extraction; see e.g. WO 2012135111 A2. The preferred solvents used
in aromatic solvent extraction are sulfolane, tetraethylene glycol
and N-methylpyrolidone which are commonly used solvents in
commercial aromatics extraction processes. These species are often
used in combination with other solvents or other chemicals
(sometimes called co-solvents) such as water and/or alcohols.
Non-nitrogen containing solvents such as sulfolane are particularly
preferred. Commercially applied dearomatization processes are less
preferred for the dearomatization of hydrocarbon mixtures having a
boiling point range that exceeds 250.degree. C., preferably
200.degree. C., as the boiling point of the solvent used in such
solvent extraction needs to be lower than the boiling point of the
aromatic compounds to be extracted. Solvent extraction of heavy
aromatics is described in the art; see e.g. U.S. Pat. No.
5,880,325. Alternatively, other known methods than solvent
extraction, such as molecular sieve separation or separation based
on boiling point, can be applied for the separation of heavy
aromatics in a dearomatization process.
A process to separate a mixed hydrocarbon stream into a stream
comprising predominantly paraffins and a second stream comprising
predominantly aromatics and naphthenes comprises processing said
mixed hydrocarbon stream in a solvent extraction unit comprising
three main hydrocarbon processing columns: solvent extraction
column, stripper column and extract column. Conventional solvents
selective for the extraction of aromatics are also selective for
dissolving light naphthenic and to a lesser extent light paraffinic
species hence the stream exiting the base of the solvent extraction
column comprises solvent together with dissolved aromatic,
naphthenic and light paraffinic species. The stream exiting the top
of the solvent extraction column (often termed the raffinate
stream) comprises the relatively insoluble, with respect to the
chosen solvent) paraffinic species. The stream exiting the base of
the solvent extraction column is then subjected, in a distillation
column, to evaporative stripping in which species are separated on
the basis of their relative volatility in the presence of the
solvent. In the presence of a solvent, light paraffinic species
have higher relative volatilities than naphthenic species and
especially aromatic species with the same number of carbon atoms,
hence the majority of light paraffinic species may be concentrated
in the overhead stream from the evaporative stripping column. This
stream may be combined with the raffinate stream from the solvent
extraction column or collected as a separate light hydrocarbon
stream. Due to their relatively low volatility the majority of the
naphthenic and especially aromatic species are retained in the
combined solvent and dissolved hydrocarbon stream exiting the base
of this column. In the final hydrocarbon processing column of the
extraction unit, the solvent is separated from the dissolved
hydrocarbon species by distillation. In this step the solvent,
which has a relatively high boiling point, is recovered as the base
stream from the column whilst the dissolved hydrocarbons,
comprising mainly aromatics and naphthenic species, are recovered
as the vapor stream exiting the top of the column. This latter
stream is often termed the extract.
The process of the present invention may require removal of sulfur
from certain crude oil fractions to prevent catalyst deactivation
in downstream refinery processes, such as catalytic reforming or
fluid catalytic cracking. Such a hydrodesulfurization process is
performed in a "HDS unit" or "hydrotreater"; see Alfke (2007) loc.
cit. Generally, the hydrodesulfurization reaction takes place in a
fixed-bed reactor at elevated temperatures of 200-425.degree. C.,
preferably of 300-400.degree. C. and elevated pressures of 1-20 MPa
gauge, preferably 1-13 MPa gauge in the presence of a catalyst
comprising elements selected from the group consisting of Ni, Mo,
Co, W and Pt, with or without promoters, supported on alumina,
wherein the catalyst is in a sulfide form.
In a further embodiment, the process further comprises a
hydrodealkylation step wherein the BTX (or only the toluene and
xylenes fraction of said BTX produced) is contacted with hydrogen
under conditions suitable to produce a hydrodealkylation product
stream comprising benzene and fuel gas.
The process step for producing benzene from BTX may include a step
wherein the benzene comprised in the hydrocracking product stream
is separated from the toluene and xylenes before hydrodealkylation.
The advantage of this separation step is that the capacity of the
hydrodealkylation reactor is increased. The benzene can be
separated from the BTX stream by conventional distillation.
Processes for hydrodealkylation of hydrocarbon mixtures comprising
C6-C9 aromatic hydrocarbons are well known in the art and include
thermal hydrodealkylation and catalytic hydrodealkylation; see e.g.
WO 2010/102712 A2. Catalytic hydrodealkylation is preferred as this
hydrodealkylation process generally has a higher selectivity
towards benzene than thermal hydrodealkylation. Preferably
catalytic hydrodealkylation is employed, wherein the
hydrodealkylation catalyst is selected from the group consisting of
supported chromium oxide catalyst, supported molybdenum oxide
catalyst, platinum on silica or alumina and platinum oxide on
silica or alumina. The process conditions useful for
hydrodealkylation, also described herein as "hydrodealkylation
conditions", can be easily determined by the person skilled in the
art. The process conditions used for thermal hydrodealkylation are
for instance described in DE 1668719 A1 and include a temperature
of 600-800.degree. C., a pressure of 3-10 MPa gauge and a reaction
time of 15-45 seconds. The process conditions used for the
preferred catalytic hydrodealkylation are described in WO
2010/102712 A2 and preferably include a temperature of
500-650.degree. C., a pressure of 3.5-8 MPa gauge, preferably of
3.5-7 MPa gauge and a Weight Hourly Space Velocity of 0.5-2 h-1.
The hydrodealkylation product stream is typically separated into a
liquid stream (containing benzene and other aromatics species) and
a gas stream (containing hydrogen, H2S, methane and other low
boiling point hydrocarbons) by a combination of cooling and
distillation. The liquid stream may be further separated, by
distillation, into a benzene stream, a C7 to C9 aromatics stream
and optionally a middle-distillate stream that is relatively rich
in aromatics. The C7 to C9 aromatic stream may be fed back to
reactor section as a recycle to increase overall conversion and
benzene yield. The aromatic stream which contains polyaromatic
species such as biphenyl, is preferably not recycled to the reactor
but may be exported as a separate product stream and recycled to
the integrated process as middle-distillate ("middle-distillate
produced by hydrodealkylation"). The gas stream contains
significant quantities of hydrogen may be recycled back the
hydrodealkylation unit via a recycle gas compressor or to any other
refinery that uses hydrogen as a feed. A recycle gas purge may be
used to control the concentrations of methane and H2S in the
reactor feed.
As used herein, the term "gas separation unit" relates to the
refinery unit that separates different compounds comprised in the
gases produced by the crude distillation unit and/or refinery
unit-derived gases. Compounds that may be separated to separate
streams in the gas separation unit comprise ethane, propane,
butanes, hydrogen and fuel gas mainly comprising methane. Any
conventional method suitable for the separation of said gases may
be employed. Accordingly, the gases may be subjected to multiple
compression stages wherein acid gases such as CO2 and H2S may be
removed between compression stages. In a following step, the gases
produced may be partially condensed over stages of a cascade
refrigeration system to about where only the hydrogen remains in
the gaseous phase. The different hydrocarbon compounds may
subsequently be separated by distillation.
A process for the conversion of alkanes to olefins involves "steam
cracking" or "pyrolysis". As used herein, the term "steam cracking"
relates to a petrochemical process in which saturated hydrocarbons
are broken down into smaller, often unsaturated, hydrocarbons such
as ethylene and propylene. In steam cracking gaseous hydrocarbon
feeds like ethane, propane and butanes, or mixtures thereof, (gas
cracking) or liquid hydrocarbon feeds like naphtha or gasoil
(liquid cracking) is diluted with steam and briefly heated in a
furnace without the presence of oxygen. Typically, the reaction
temperature is 750-900.degree. C., but the reaction is only allowed
to take place very briefly, usually with residence times of 50-1000
milliseconds. Preferably, a relatively low process pressure is to
be selected of atmospheric up to 175 kPa gauge. Preferably, the
hydrocarbon compounds ethane, propane and butanes are separately
cracked in accordingly specialized furnaces to ensure cracking at
optimal conditions. After the cracking temperature has been
reached, the gas is quickly quenched to stop the reaction in a
transfer line heat exchanger or inside a quenching header using
quench oil. Steam cracking results in the slow deposition of coke,
a form of carbon, on the reactor walls. Decoking requires the
furnace to be isolated from the process and then a flow of steam or
a steam/air mixture is passed through the furnace coils. This
converts the hard solid carbon layer to carbon monoxide and carbon
dioxide. Once this reaction is complete, the furnace is returned to
service. The products produced by steam cracking depend on the
composition of the feed, the hydrocarbon to steam ratio and on the
cracking temperature and furnace residence time. Light hydrocarbon
feeds such as ethane, propane, butane or light naphtha give product
streams rich in the lighter polymer grade olefins, including
ethylene, propylene, and butadiene. Heavier hydrocarbon (full range
and heavy naphtha and gas oil fractions) also give products rich in
aromatic hydrocarbons.
To separate the different hydrocarbon compounds produced by steam
cracking the cracked gas is subjected to a fractionation unit. Such
fractionation units are well known in the art and may comprise a
so-called gasoline fractionator where the heavy-distillate ("carbon
black oil") and the middle-distillate ("cracked distillate") are
separated from the light-distillate and the gases. In the
subsequent optional quench tower, most of the light-distillate
produced by steam cracking ("pyrolysis gasoline" or "pygas") may be
separated from the gases by condensing the light-distillate.
Subsequently, the gases may be subjected to multiple compression
stages wherein the remainder of the light distillate may be
separated from the gases between the compression stages. Also acid
gases (CO2 and H2S) may be removed between compression stages. In a
following step, the gases produced by pyrolysis may be partially
condensed over stages of a cascade refrigeration system to about
where only the hydrogen remains in the gaseous phase. The different
hydrocarbon compounds may subsequently be separated by simple
distillation, wherein the ethylene, propylene and C4 olefins are
the most important high-value chemicals produced by steam cracking.
The methane produced by steam cracking is generally used as fuel
gas, the hydrogen may be separated and recycled to processes that
consume hydrogen, such as hydrocracking processes. The acetylene
produced by steam cracking preferably is selectively hydrogenated
to ethylene. The alkanes comprised in the cracked gas may be
recycled to the process for olefins synthesis.
The term "propane dehydrogenation unit" as used herein relates to a
petrochemical process unit wherein a propane feedstream is
converted into a product comprising propylene and hydrogen.
Accordingly, the term "butane dehydrogenation unit" relates to a
process unit for converting a butane feedstream into C4 olefins.
Together, processes for the dehydrogenation of lower alkanes such
as propane and butanes are described as lower alkane
dehydrogenation process. Processes for the dehydrogenation of lower
alkanes are well-known in the art and include oxidative
dehydrogenation processes and non-oxidative dehydrogenation
processes. In an oxidative dehydrogenation process, the process
heat is provided by partial oxidation of the lower alkane(s) in the
feed. In a non-oxidative dehydrogenation process, which is
preferred in the context of the present invention, the process heat
for the endothermic dehydrogenation reaction is provided by
external heat sources such as hot flue gases obtained by burning of
fuel gas or steam. In a non-oxidative dehydrogenation process the
process conditions generally comprise a temperature of
540-700.degree. C. and an absolute pressure of 25-500 kPa. For
instance, the UOP Oleflex process allows for the dehydrogenation of
propane to form propylene and of (iso)butane to form (iso)butylene
(or mixtures thereof) in the presence of a catalyst containing
platinum supported on alumina in a moving bed reactor; see e.g.
U.S. Pat. No. 4,827,072. The Uhde STAR process allows for the
dehydrogenation of propane to form propylene or of butane to form
butylene in the presence of a promoted platinum catalyst supported
on a zinc-alumina spinel; see e.g. U.S. Pat. No. 4,926,005. The
STAR process has been recently improved by applying the principle
of oxydehydrogenation. In a secondary adiabatic zone in the reactor
part of the hydrogen from the intermediate product is selectively
converted with added oxygen to form water. This shifts the
thermodynamic equilibrium to higher conversion and achieves a
higher yield. Also the external heat required for the endothermic
dehydrogenation reaction is partly supplied by the exothermic
hydrogen conversion. The Lummus Catofin process employs a number of
fixed bed reactors operating on a cyclical basis. The catalyst is
activated alumina impregnated with 18-20 wt-% chromium; see e.g. EP
0 192 059 A1 and GB 2 162 082 A. The Catofin process has the
advantage that it is robust and capable of handling impurities
which would poison a platinum catalyst. The products produced by a
butane dehydrogenation process depends on the nature of the butane
feed and the butane dehydrogenation process used. Also the Catofin
process allows for the dehydrogenation of butane to form butylene;
see e.g. U.S. Pat. No. 7,622,623.
The present invention will be discussed in the next Example which
example should not be interpreted as limiting the scope of
protection.
The sole FIGURE provides a schematic flow sheet of an embodiment of
the present invention.
EXAMPLE
The process scheme can be found in the sole FIGURE. A hydrocarbon
feedstock 29 is fed into a reaction area for ringopening 4 and its
reaction products, which are generated from said reaction area, are
separated into an overhead stream 9 and a side stream 13. The side
stream 13 is fed into a gasoline hydrocracker (GHC) unit 5, wherein
the reaction products of said GHC unit 5 are separated into an
overhead gas stream 33, comprising light components such C2-C4
paraffins, hydrogen and methane, and a stream 15 comprising
predominantly aromatic hydrocarbon compounds and non-aromatic
hydrocarbon compounds. The overhead gas stream 33 from the gasoline
hydrocracker (GHC) unit 5 is fed as feedstock 8 to a steam cracker
unit 1.
In a preferred embodiment hydrocarbon feedstock 7 can be divided in
a feed 28 and a feed 12, wherein feed 28 is pretreated in an
aromatics extraction unit 3. From aromatics extraction unit 3 its
bottom stream 34 is fed into said reaction area for ringopening 4
and its overhead stream 26 is fed into said steam cracker unit 1.
The aromatics extraction unit 3 is chosen from the group of the
type of a distillation unit, a solvent extraction unit and a
molecular sieve, or even a combination thereof. For example light
Naphtha 6 is a feedstock directly sent to the steam cracker unit
1.
In a preferred embodiment the C2-C4 paraffins are separated from
said overhead gas stream 33, and the C2-C4 paraffins thus separated
are sent to the furnace section of steam cracker unit 1. In another
preferred embodiment the C2-C4 paraffins are separated in
individual streams, each stream predominantly comprising C2
paraffins, C3 paraffins and C4 paraffins, respectively, and each
individual stream is fed to a specific furnace section of steam
cracker unit 1. Such a separation of C2-C4 paraffins from said
overhead gas stream 33 is carried out by cryogenic distillation or
solvent extraction.
The reaction products 18 of said steam cracking unit 1 are
separated in separator 2 into an overhead stream 17, comprising
C2-C6 alkanes, a middle stream 14, which contains C2-olefins,
C3-olefins and C4-olefins, and a first bottom stream 19 comprising
C9+ hydrocarbons, and a second bottom stream 10 comprising aromatic
hydrocarbon compounds and non-aromatic hydrocarbon compounds.
Second bottom stream 10 comprises pygas. Hydrogen and methane can
be recovered from separator 2 as well and re-used elsewhere. The
overhead stream 17 is returned to said steam cracking unit 1.
Second bottom stream 10 is fed into said gasoline hydrocracker
(GHC) unit 5. First bottom stream 19 predominantly containing
carbon black oil (CBO) and cracked distillates (CD) is fed into
said reaction area for ringopening 4.
In a preferred embodiment stream 15 from said gasoline hydrocracker
(GHC) unit 5 is further separated in a BTX rich fraction and in
heavy fraction (not shown). Overhead stream 33 from the gasoline
hydrocracker (GHC) unit 5 is divided into a stream 8 and a stream
20, wherein stream 20 is fed to a dehydrogenation unit 23. As
mentioned before, it is preferred to send only the C3-C4 fraction
of overhead stream 33 to the dehydrogenation unit 23. Overhead
stream 9 from the reaction area for ring opening 4 can also be fed
into dehydrogenation unit 23 and/or into steam cracker unit 1. And
for this stream 9 it is also preferred to send only the C3-C4
fraction of stream 9 to the dehydrogenation unit 23. According to a
preferred embodiment the C3-C4 fractions are recovered from both
stream 9 and stream 33 in a single process unit and these C3-C4
fractions are sent to the dehydrogenation unit 23. This means that
after suitable processing hydrogen and methane are removed from
stream 9 and stream 33 before sending stream 20 to the
dehydrogenation unit 23.
In a preferred embodiment a stream 25 rich in mono aromatics is
recovered from said hydrocarbon feedstock 24 and stream 25 thus
recovered is directly fed into said gasoline hydrocracker (GHC)
unit 5. The remaining part 32 of feedstock 24 is sent to a reaction
area for ringopening 4.
Hydrogen 27 can be recovered from the reaction products 18 of said
steam cracking unit 1 and the hydrogen 27 thus recovered can be
sent to said gasoline hydrocracker (GHC) unit 5 and/or said
reaction area for ring opening 4 via line 22 and line 31,
respectively. In another embodiment it is also possible to recover
hydrogen 21 from said dehydrogenation unit 23 and the hydrogen 21
thus recovered can be fed into said gasoline hydrocracker (GHC)
unit 5 and/or said reaction area for ring opening 4.
According to the process scheme of FIGURE feedstock 7 can be
divided in a feedstock 28 and a feedstock 12, wherein feedstock 12
does not undergo an extraction in the aromatics extraction unit 3.
Feedstock 12 can be mixed with other types of feedstock 29, if
appropriate, and the combined feedstock 16, after being mixed, if
necessary, with the bottom stream 34 of unit 3, is now indicated as
reference number 24. In a preferred embodiment mono aromatics 25
are separated from feedstock 24 in unit 30 and the stream 32 thus
obtained is fed into unit 4.
The Example disclosed herein makes a distinction between several
cases.
According to case 1 kerosine as feedstock is sent directly to steam
cracker unit (comparative example).
According to case 2 (example according to the invention) kerosine
as feedstock is sent to a reaction area for ringopening and the
side stream thereof is sent to a gasoline hydrocracker (GHC) unit,
the LPG fraction from GHC being steam cracked.
According to case 3 (example according to the invention) kerosine
as feedstock is first pretreated in an aromatics extraction unit,
wherein the paraffins fraction is sent to a steam cracker unit and
the naphthenes and aromatics fraction is sent to a reaction area
for ringopening and the side stream thereof is sent to a gasoline
hydrocracker (GHC) unit, the LPG fraction from GHC being steam
cracked.
Case 4 (example according to the present invention) is similar to
case 2 but the feedstock in case 4 is now LVGO.
The characteristics of kerosine and LVGO can be found in Table
1.
TABLE-US-00001 TABLE 1 characteristics of kerosine and LVGO
Kerosine LVGO n-Paraffins wt-% 23.7 18.3 i-Paraffins wt-% 17.9 13.8
Naphthenes wt-% 37.4 35.8 Aromatics wt-% 21.0 32.0 Density 60 F.
kg/L 0.810 0.913 IBP .degree. C. 174 306 BP10 .degree. C. 196 345
BP30 .degree. C. 206 367 BP50 .degree. C. 216 384 BP70 .degree. C.
226 404 BP90 .degree. C. 242 441 FBP .degree. C. 266 493
The conditions of the steam cracker unit are as follows: ethane and
propane furnaces:coil outlet temperature=845.degree. C.,
steam-to-oil-ratio=0.37, C4-furnaces: coil outlet
temperature=820.degree. C., Steam-to-oil-ratio=0.37, liquid
furnaces: coil outlet temperature=820.degree. C.,
steam-to-oil-ratio=0.37.
Table 2 shows the battery limit product slate (wt. % of
feedstock).
TABLE-US-00002 TABLE 2 the battery limit product slate (wt. % of
feedstock) CASE 2 CASE 3 KEROSINE TO KEROSINE TO CASE 4 PARTIAL
DEARO, paraffins to LVGO TO PARTIAL CASE 1 RINGOPENING + SC, arom +
naphthenes RINGOPENING + BATTERY LIMIT PRODUCT SLATE KEROSINE to SC
GHC + SC to PARO - GHC GHC + SC H2 production (SC) 0.6 2.5 1.6 2.5
H2 consumption (P-ARO + GHC) 0 3.5 2.1 4.5 CH4 14.4 18.6 15.8 18.9
ETHYLENE 29.0 47.7 42.6 48.6 PROPYLENE 15.1 12.4 15.4 12.6
BUTADIENE 4.9 2.4 4.1 2.5 ISO-BUTENE 2.0 0.5 0.9 0.5 BENZENE 7.9
5.4 7.2 5.0 TX CUT 4.0 6.4 7.1 5.6 STYRENE 1.6 3.0 3.2 2.6 OTHER
C7-C8 2.3 0.3 0.6 0.3 C9 RESIN FEED 4.8 0.1 0.2 0.1 CD 1.6 0.0 0.2
0.0 CBO 11.6 0.5 1.0 0.6 % HIGH VALUE CHEMICALS 66.8 78.1 81.1
77.7
For each case the hydrogen balance was calculated. For case 1 the
H2 balance is +0.6%, for case 2 the H2 balance is -1.0%, for case 3
the H2 balance is -0.5%, and for case 4 the H2 balance is -2.0%,
respectively.
The data presented above show that the presence of a reaction area
for ringopening and gasoline hydrocracking (GHC) of the diesel
converts aromatics into BTX and LPG and converts naphthenes into
LPG. The steam cracker product from this LPG contains increased
olefins yields (ethylene and propylene), increased CH4 yield and
decreased C9+ yield (compared to steam cracking diesel straight
away as in case 1). The present inventors found that this effect
also applies to LVGO and HVGO. It is to be noted that a reaction
area for ringopening requires additional H.sub.2, i.e. a negative
hydrogen balance for cases 2, 3, and 4. Moreover, when applying the
propane dehydrogenation (PDH)/butane dehydrogenation (BDH) options
a positive hydrogen balance can be achieved. Furthermore, the rise
in ethylene is also highly remarkable in the method according to
the present invention.
* * * * *